Process for hydrogenating a coke-forming hydrocarbon distillate

ABSTRACT

A process for hydrotreating (hydroprocessing) hydrocarbons and mixtures of hydrocarbons utilizing a catalytic composite comprising a combination of a nickel component and a tungsten component with a silica-alumina carrier material wherein said carrier material is co-gelled silica-alumina consisting of from about 43 percent to about 57 percent by weight of alumina and from about 57 percent to 43 percent by weight silica and wherein said components are present in amounts sufficient to result in the composite containing, on an elemental basis, about 2 to about 10 percent by weight of the nickel component and about 8 to about 20 percent by weight of the tungsten component, in which process there is effected a chemical consumption of hydrogen. Key features of the subject composite are the criticality of the alumina content of the carrier material and the facility of using a co-gelled silica-alumina carrier material. The principal utility of the subject composite is in the hydrocracking of hydrocarbons. A specific example of the catalyst disclosed is a combination of nickel and tungsten with a co-gelled silica-alumina carrier material containing 50 weight percent alumina in amounts sufficient to result in the composite containing, on an elemental basis, about 7 to about 9 weight percent nickel and about 17 to about 19 weight percent tungsten. Other hydrocarbon hydroprocesses are directed toward the hydrogenation of aromatic nuclei, the ring-opening of cyclic hydrocarbons, desulfurization, denitrification, hydrogenation, etc.

CROSS-REFERENCE TO RELATED APPLICATIONS

The present application is a Division of my copending application, Ser.No. 716,922, filed Aug. 23, 1976, now U.S. Pat. No. 4,061,563, which isa Division of application, Ser. No. 599,561, filed July 28, 1975, nowU.S. Pat. No. 4,003,956, issued Jan. 18, 1977 which is acontinuation-in-part of application, Ser. No. 466,759, filed May 3,1974, now U.S. Pat. No. 3,931,048, all the teachings of whichapplications are incorporated by specific reference thereto.

APPLICABILITY OF INVENTION

The present invention encompasses the use of a catalytic compositecomprising a combination operating a nickel component and a tungstencomponent with a silica-alumina carrier material wherein said carriermaterial is co-gelled silica-alumina consisting of from about 43 percentto about 57 percent by weight of alumina and from about 57 percent to 43percent by weight silica and wherein said components are present inamounts sufficient to result in the composite containing on an elementalbasis, about 2 to about 10 percent by weight of the nickel component andabout 8 to about 20 percent by weight of the tungsten component in thehydrotreating of hydrocarbons and mixtures of hydrocarbons. As utilizedherein, the term "hydrotreating" is intended to be synonymous with theterm "hydroprocessing", which involves the conversion of hydrocarbons atopening conditions selected to effect a chemical consumption ofhydrogen. Included within the processes intended to be encompassed bythe term "hydroprocessing" are hydrocracking, aromatic hydrogenation,ring-opening, hydrorefining (for nitrogen removal and olefinsaturation), desulfurization (often included in hydrorefining) andhydrogenation, etc. As will be recognized, one common attribute of theseprocesses, and the reactions being effected therein, is that they areall "hydrogen-consuming", and are, therefore, exothermic in nature.

The individual characteristics of the foregoing hydrotreating processes,including preferred operating conditions and techniques, will behereinafter described in greater detail. The subject of the presentinvention is the use of a catalytic composite which has exceptionalactivity and resistance to deactivation when employed in ahydrogen-consuming process. Such processes require a catalyst havingboth a hydrogenation function and a cracking function. Morespecifically, the present process uses a dual-function catalyticcomposite which enables substantial improvements in those hydroprocessesthat have traditionally used a dual-function catalyst. The particularcatalytic composite comprising a combination of a nickel component and atungsten component with a silica-alumina carrier material wherein saidcarrier material is co-gelled silica-alumina consisting of from about 43percent to about 57 percent by weight of alumina and from about 57percent to 43 percent by weight silica and wherein said components arepresent in amounts sufficient to result in the composite containing, onan elemental basis, about 2 to about 10 percent by weight of the nickelcomponent and about 8 to about 20 percent by weight of the tungstencomponent; specifically, an improved hydrocracking process utilizes acatalytic composite as hereinabove described for improved activity,product selectivity and operational stability characteristics.

Composites having dual-function catalytic activity are widely employedin many industries for the purpose of accelerating a wide spectrum ofhydrocarbon conversion reactions. Generally, the cracking function isthought to be associated with an acid-acting material of the porous,adsorptive refractory inorganic oxide type which is typically utilizedas the carrier material for a metallic component from the metals, orcompounds of metals, of Groups V through VIII of the Periodic Table, towhich the hydrogenation function is generally attributed.

Catalytic composites are used to promote a wide variety of hydrocarbonconversion reactions such as hydrocracking, isomerization,dehydrogenation, hydrogenation, desulfurization, reforming,ring-opening, cyclization, aromatization, alkylation andtransalkylation, polymerization, cracking, etc., some of which reactionsare hydrogen-producing producing while others are hydrogen-consuming. Inusing the term "hydrogen-consuming", I intend to exclude those processeswherein the only hydrogen consumption involves the saturation of lightolefins, resulting from undesirable cracking, which produces the lightparaffins, methane, ethane and propane. It is to the latter group ofreactions, hydrogen-consuming, that the present invention is applicable.In many instances, the commercial application of these catalysts is inprocesses where more than one of these reactions proceed simultaneously.An example of this type of process is a hydrocracking process whereincatalysts are utilized to effect selective hydrogenation and cracking ofhigh molecular weight materials to produce a lower-boiling, morevaluable output stream. Another such example would be the conversion ofaromatic hydrocarbons into jet fuel components, principally straight, orslightly branched paraffins.

Regardless of the reaction involved, or the particular process, it isvery important that the catalyst exhibit not only the capability toperform its specified functions initially, but also perform themsatisfactorily for prolonged periods of time. The analytical termsemployed in the art to measure how efficient a particular catalystperforms its intended functions in a particular hydrocarbon conversionprocess, are activity, selectivity and stability. For the purpose ofdiscussion, these terms are conveniently defined herein, for a givencharge stock, as follows: (1) activity is a measure of the ability ofthe catalyst to convert a hydrocarbon feed stock into products at aspecified severity level, where severity level alludes to the operatingconditions employed -- the temperature, pressure, liquid hourly spacevelocity and hydrogen concentration; (2) selectivity refers to theweight percent or volume percent of the reactants that are convertedinto the desired product and/or products; (3) stability connotes therate of change of the activity and selectivity parameters with time --obviously, the smaller rate implying the more stable catalyst. Withrespect to a hydrogen-consuming process, for example hydrocracking,activity, stability and selectivity are similarly defined. Thus,"activity" connotes the quantity of charge stock, boiling above a giventemperature, which is converted to hydrocarbons boiling below the giventemperature. "Selectivity" refers to the quantity of converted chargestock which boils below the desired end point of the product, as well asabove a minimum specified initial boiling point. "Stability" connotesthe rate of change of activity and selectivity. Thus, for example, wherea gas oil, boiling above about 650° F., is subjected to hydrocracking,"activity" connotes the conversion of 650° F.-plus charge stock to 650°F.-minus product. "Selectivity" can allude to the quantity of conversioninto gasoline boiling range hydrocarbons -- i.e., pentanes and heavier,normally liquid hydrocarbons boiling up to about 400° F. " Stability"might be conveniently expressed in terms of temperature increaserequired during various increments of catalyst life, in order tomaintain the desired activity.

As is well known to those skilled in the art, the principal cause ofobserved deactivation or instability of a dual-function catalyst isassociated with the fact that coke forms on the surface of the catalystduring the course of the reaction. More specifically, in the varioushydrocarbon conversion processes, and especially those which arecategorized as hydrogen-consuming, the operating conditions utilizedresult in the formation of high molecular weight, black solid orsemi-solid hydrogen-poor carbonaceous material which coats the surfaceof the catalyst and reduces its activity by shielding its active sitesfrom the reactants. Accordingly, a major problem facing workers in thisarea is the development of more active and selective catalyticcomposites that are not as sensitive to the presence of thesecarbonaceous materials and/or have the capability to suppress the rateof formation of these materials at the operating conditions employed ina particular process.

I have now found a dual-function catalyst composite which possessesimproved activity, selectivity and stability when employed in thehydroprocessing of hydrocarbons, wherein there is effected a chemicalconsumption of hydrogen. In particular, I have found that the use of acatalytic composite comprising a combination of a nickel component and atungsten component with a silica-alumina carrier material wherein saidcarrier material is co-gelled silica-alumina consisting of from about 43percent to about 57 percent by weight of alumina and from about 57percent to 43 percent by weight silica and wherein said components arepresent in amounts sufficient to result in the composite containing, onan elemental basis, about 2 to about 10 percent by weight of the nickelcomponent and about 8 to about 20 percent by weight of the tungstencomponent improves the overall operation of these hydrogen-consumingprocesses. Moreover, I have determined that a catalytic composite ashereinabove described, when utilized in a process for hydrocrackinghydrocarbonaceous material into lower-boiling hydrocarbon products,affords substantial improvement in performance and results. Asindicated, the present invention essentially enables the performancecharacteristics of the process to be sharply and materially improved.

Destructive hydrogenation by catalytic means, more commonly calledhydrocracking, is old and well-known to the art. Destructivehydrogenation of the hydrocarbon oil, which is usually a coal tar or ahigh-boiling petroleum fraction, such as gas oils or topped crude,generally is performed at relatively high temperatures and pressures ofthe order of 750° F. and 1500 psig., and upward. Catalysts for thedestructive hydrogenation of hydrocarbons are generally a combination ofhydrogenation and cracking catalysts.

While many types of catalyst compositions have been proposed fordestructive hydrogenation or hydrocracking, it has been found thatcatalysts comprised of silica, alumina, tungsten and nickel areespecially suitable. Such catalysts are well known in the hydrocrackingart.

From U.S. Pat. No. 3,216,922, a process is known for the preparation ofhydrocracking catalysts comprising a silica-alumina mixture as a carrierin which the carrier is obtained by first precipitating silica gel froma water glass solution and subsequently, after aging of the gel,precipitating aluminum hydroxide thereon. As the aluminum salt fromwhich the aluminum hydroxide is formed, use is made of aluminum sulphatewhich is added in such a quantity that the molar ratio of silica toalumina in the finished carrier is approximately 5:1. It was found,however, that the use of hydrocracking catalysts, of which the carrierwas obtained in the manner described, produced less favorable results inthe hydrocracking of flashed distillates.

In an effort to prepare a more satisfactory hydrocracking catalyst,British Pat. No. 1,183,778 teaches a process for the preparation of analumina-silica-nickel-tungsten hydrocracking catalyst which comprisespreparing a catalyst carrier by first precipitating from an aqueoussolution comprising silicate ions, a silica gel, subjecting the gel toaging at elevated temperature, precipitating aluminum hydroxide on theaged gel by addition of an aqueous aluminum nitrate solution and analkaline-reacting solution, separating, drying and finally calcining theresulting precipitate of aluminum hydroxide on silica gel and thensupporting tungsten and nickel on the catalyst carrier and subsequentlyoxidizing the carrier comprising the metal salts.

However, because commercial scale hydrocracking of hydrocarbons isperformed at low space velocities, catalyst cost is an appreciablefactor in both the initial investment and operating costs ofhydrocracking plants. For this reason, there is considerable incentiveto manufacture such catalysts by the most economic method whileimproving the catalyst activity. I have discovered an improved processfor the preparation of tungsten-nickel on silica-alumina hydrocrackingcatalyst.

More specifically, I have found that a co-gelled silica-alumina is anexceptionally suitable carrier material. A co-gelled silica-aluminacarrier material in addition to being catalytically suitable is moresimply and economically produced than the prior art carriers. Theproduction of prior art carriers has been a multi-step process which hasrequired the expenditure of time and effort far in excess of thatrequired for a co-gelled carrier.

More specifically, my process is an improved process for the preparationof such catalyst wherein the nickel component is present in an amountfrom about 2 weight percent to about 10 weight percent, and the tungstencomponent is present in an amount from about 8 weight percent to 20weight percent. I have also discovered that an unusually superiorcatalyst results if the catalyst base contains from about 43 percent toabout 57 percent by weight of alumina. The criticality of the range ofthe alumina concentration is further illustrated hereinbelow.

A particularly preferred co-gelled silica-alumina catalyst basecomprises from about 43 percent to about 57 percent alumina and fromabout 57 percent to about 43 percent silica and still more preferablyfrom about 48 percent to about 52 percent alumina.

In addition to the foregoing compositional limitations, it is importantthat the catalyst base have adequate pore volume, that is, a pore volumeof at least 0.5 cc./g. and preferably at least 0.6 cc./g. or even 0.7cc./g.

The co-gelled silica-alumina catalyst base is preferably in the xerogelstate, that is, it is dried sufficiently to afford the usual microporousstructure and therefore an appreciable available surface. It is alsopossible to use a rigid silica-alumina catalyst base which has merelybeen dried at a relatively low temperature, e.g., 125° C., and whichstill contains considerable amounts of water. In this latter case,however, the degree of drying must nevertheless be sufficient to removeessentially all water from the pores of the base.

The catalyst of the present invention can be utilized to achieve themaximum production of LPG (liquefied petroleum gas) in thepropane/butane range from naphtha or gasoline boiling range distillates.Heavier charge stocks, including kerosenes, light gas oils, heavy gasoils, full boiling range gas oils and "black oils" may be readilyconverted into lower-boiling, normally liquid products includinggasolines, kerosenes, middle-distillates, lube oils, etc.

In one embodiment, accordingly, the present invention provides a methodof preparing catalysts having hydrocracking activity comprising thesteps: (a) co-gelling a silica-alumina carrier material consisting offrom about 43 percent to about 57 percent by weight alumina and fromabout 57 to 43 percent by weight silica; (b) impregnating saidsilica-alumina carrier material with an aqueous solution of a nickelsalt and a tungsten salt, the concentrations of the salts in the aqueoussolution being sufficient to deposit on the carrier material an amountof salts equivalent to 2 to 10 percent by weight nickel and 8 to 20percent by weight tungsten based on the total weight of the finishedcatalyst; and, (c) calcining the impregnated carrier material.

In a second embodiment, the present invention relates to a process forhydrocracking hydrocarbons which process comprises reacting saidhydrocarbons with hydrogen in a reaction zone containing a catalyticcomposite prepared by a method comprising the steps: (a) co-gelling asilica-alumina carrier material consisting of from about 43 percent toabout 57 percent by weight alumina and from about 57 to 43 percent byweight silica; (b) impregnating said silica-alumina carrier materialwith an aqueous solution of a nickel salt and a tungsten salt, theconcentrations of the salts in the aqueous solution being sufficient todeposit on the carrier material an amount of salts equivalent to 2 to 10percent by weight nickel and 8 to 20 percent by weight tungsten based onthe total weight of the finished catalyst; and, (c) calcining theimpregnated carrier material.

In a specific embodiment, the hydrocracking conditions include a maximumcatalyst bed temperature of about 600° to about 900° F., a pressure ofabout 500 to about 5000 psig., a liquid hourly space velocity of about0.1 to about 10 and a hydrogen circulation rate in the range of about1,000 to about 50,000 scf./bbl.

In another specific embodiment, the catalytic composite is oxidized, inan atmosphere of air, at a temperature above 1000° F. prior to contactwith the fresh feed charge stock.

Another embodiment relates to a catalytic composite, comprising acombination of a nickel component and a tungsten component with asilica-alumina carrier material wherein said carrier material isco-gelled silica-alumina consisting of from about 43 percent to about 57percent by weight alumina and from about 57 percent to 43 percent byweight silica and wherein said components are present in amountssufficient to result in the composite containing, on an elemental basis,about 2 to about 10 percent by weight of the nickel component and about8 to about 20 percent by weight of the tungsten component.

An object of the present invention is to afford a process for thehydroprocessing of a hydrocarbon, or mixture of hydrocarbons. Acorollary objective is to improve the selectivity and stability ofhydroprocessing utilizing a highly active catalytic composite of thepresent invention.

A specific object of my invention resides in the improvement ofhydrogen-consuming processes including hydrocracking, hydrorefining,ring-opening for jet fuel production, hydrogenation or aromatichydrocarbons, desulfurization, denitrification, etc. Therefore, in oneembodiment, the present invention encompasses a hydrocarbon hydroprocesswhich comprises reacting a hydrocarbon with hydrogen at conditionsselected to effect chemical consumption of hydrogen and in contact witha catalytic composite as hereinabove described. In another embodiment,the operating conditions include a pressure of from 400 to about 5,000psig., an LHSV (defined as volumes of liquid hydrocarbon charge per hourper volume of catalyst disposed in the reaction zone) of from 0.1 toabout 10.0, a hydrogen circulation rate of from 1,000 to about 50,000scf./Bbl. and a maximum catalyst temperature of from 200° to about 900°F.

In another embodiment, the process is further characterized in that thecatalytic composite is reduced and sulfided prior to contacting thehydrocarbon feed stream. In still another embodiment, my inventioninvolves a process for hydrogenating a coke-forming hydrocarbondistillate containing di-olefinic and mono-olefinic hydrocarbons, andaromatics, which process comprises reacting said distillate withhydrogen, at a temperature below about 500° F., in contact with acatalytic composite comprising a combination of a nickel component and atungsten component with a silica-alumina carrier material wherein saidcarrier material is co-gelled silica-alumina consisting of from about 43percent to about 57 percent by weight of alumina and from about 57percent to 43 percent by weight silica and wherein said components arepresent in amounts sufficient to result in the composite containing, onan elemental basis, about 2 to about 10 percent by weight of the nickelcomponent and about 8 to about 20 percent by weight of the tungstencomponent and an alkali metal component, and recovering anaromatic/mono-olefinic hydrocarbon concentrate substantially free fromconjugated di-olefinic hydrocarbons.

Another embodiment affords a catalytic composite comprising asubstantially pure crystalline aluminosilicate material, at least about90.0% by weight of which is zeolitic.

Other objects and embodiments of my invention related to additionaldetails regarding preferred catalytic ingredients, the concentration ofcomponents in the catalytic composite, methods of catalyst preparation,individual operating conditions for use in the various hydrotreatingprocesses, preferred processing techniques and the like particularswhich are hereinafter given in the following more detailed summary of myinvention.

As hereinabove set forth, the present invention involves thehydroprocessing of hydrocarbons and mixtures of hydrocarbons, utilizinga particular catalytic composite. This catalyst comprises a combinationof a nickel component and a tungsten component with a silica-aluminacarrier material wherein said carrier material is co-gelledsilica-alumina consisting of from about 43 percent to about 57 percentby weight of alumina and from about 57 percent to 43 percent by weightsilica and wherein said components are present in amounts sufficient toresult in the composite containing, on an elemental basis, about 2 toabout 10 percent by weight of the nickel component and about 8 to about20 percent by weight of the tungsten component.

Catalytic composites, tailored for the conversion of hydrocarbonaceousmaterial and particularly those intended for utilization in ahydrocracking process, have traditionally consisted of metallic elementschosen from Group VIII of the Periodic Table; however, metalliccomponents from Group VI-B are quite often incorporated therein. Inthose instances where hydrocracking is intended to be accompanied bysome hydrorefining (desulfurization, denitrification, etc.) thepreferred metallic components have been nickel and molybdenum, andnickel and tungsten, which components are usually combined with a porouscarrier material comprising both alumina and silica, either amorphous orzeolitic in nature. Ample evidence may be found in the literature whichconfirms the ability of the nickel component to effect both cracking andhydrogenation reactions. Furthermore, the prior art indicates apreference for two particular methods of catalyst preparation.Predominantly preferred is an impregnating procedure wherein apreviously calcined, preformed carrier material, which is precipitatedin a multi-step manner as hereinabove described, is contacted withsuitable soluble compounds of the nickel component and the Group VI-Bcomponent, where the latter is utilized. Impregnation involves bothsubsequent drying at a temperature of about 300° F., and oxidation at atemperature of about 1100° F. The second preferred preparation schemeinvolves coprecipitating all the catalyst components, including those ofthe carrier material.

I have found that a particularly effectivesilica-alumina-nickel-tungsten hydrocracking catalyst can be preparedwhen the alumina content of the co-gelled silica-alumina support ismaintained within the critical range of from about 43 percent to about57 percent by weight alumina. Thus, it is now possible to prepare a moreactive and stable hydrocracking catalyst.

As is customary in the art of catalysis, when referring to thecatalytically active metal, or metals, it is intended to encompass theexistence of such metal in the elemental state or in some form such asan oxide, sulfide, halide, etc. Regardless of the state in which themetallic components actually exist, the concentrations are computed asif they existed in the elemental state.

The co-gelled silica-alumina may be prepared and utilized as spheres,pills, pellets, extrudates, granules, etc. In a preferred method ofmanufacture, an aqueous water glass solution, diluted to a silicaconcentration of from about 5 to about 15 weight percent, is acidifiedwith hydrochloric acid or other suitable mineral acid. The resulting solis acid aged at a pH of from about 4 to about 4.8 to form a hydrogel,and the hydrogel is further aged at a pH of from about 6.5 to about 7.5.The silica hydrogel is then thoroughly admixed with an aqueous aluminumsalt solution of sufficient concentration to provide a desirable aluminacontent in the silica-alumina product. The silica-alumina sol is thenprecipitated at a pH of about 8 by the addition of a basic precipitatingagent, suitably aqueous ammonium hydroxide. The silica-alumina, whichexists as a hydrogel slurried in a mother liquor, is recovered byfiltration, water-washed and dried at a temperature of from about 200°to about 500° F. Drying is preferably by spray-drying techniques wherebythe co-gelled silica-alumina is recovered as microspheres, admixed witha suitable binding agent, such as graphite, polyvinyl alcohol, etc., andextruded or otherwise compressed into pills or pellets or uniform sizeand shape.

The particularly preferred method for preparing the co-gelledsilica-alumina support is by the well known oil-drop method whichpermits the utilization of the support in the form of macrospheres. Forexample, an alumina sol, utilized as an alumina source, is commingledwith an acidified water glass solution as a silica source, and themixture further commingled with a suitable gelling agent, for example,urea, hexamethylenetetramine, or mixtures thereof. The mixture isdischarged while still below gellation temperature, and by means of anozzle or rotating disk, into a hot oil bath maintained at gellationtemperature. The mixture is dispersed into the oil bath as dropletswhich form into spheroidal gel particles during passage therethrough.The alumina sol is preferably prepared by a method wherein aluminumpellets are commingled with a quantity of treated or deionized water,with hydrochloric acid being added thereto in a sufficient amount todigest a portion of the aluminum metal and form the desired sol. Asuitable reaction rate is effected at about reflux temperature of themixture.

The spheroidal gel particles prepared by the oil-drop method are aged,usually in the oil bath, for a period of at least 10-16 hours, and thenin a suitable alkaline or basic medium for at least 3 to about 10 hours,and finally water-washed. Proper gellation of the mixture in the oilbath, as well as subsequent aging of the gel spheres, is not readilyaccomplished below about 120° F., and at about 210° F., the rapidevolution of the gases tends to rupture and otherwise weaken thespheres. By maintaining sufficient superatmospheric pressure during theforming and aging steps in order to maintain water in the liquid phase,a higher temperature can be employed, frequently with improved results.If the gel particles are aged at superatmospheric pressure, no alkalineaging step is required.

The spheres are water-washed, preferably with water containing a smallamount of ammonium hydroxide and/or ammonium nitrate. After washing, thespheres are dried, at a temperature of from about 200° to about 600° F.for a period of from about 6 to about 24 hours or more, and thencalcined at a temperature of from about 800° to about 1400° F. for aperiod of from about 2 to about 12 hours or more.

The nickel component and the tungsten component are composited with theco-gelled silica-alumina carrier material by any suitableco-impregnation technique. Thus, the carrier material can be soaked,dipped, suspended, or otherwise immersed in an aqueous impregnatingsolution containing a soluble nickle salt and a soluble tungsten salt.One suitable method comprises immersing the carrier material in theimpregnating solution and evaporating the same to dryness in a rotarysteam dryer, the concentration of the impregnating solution being suchas to ensure a final catalyst composite comprising from about 2 to about10 percent by weight nickel and 8 to about 20 percent by weighttungsten.

The catalyst composite is usually dried at a temperature of from about200° to about 500° F. for a period of from about 1 to about 10 hoursprior to calcination. In accordance with the present invention,calcination is effected in an oxidizing atmosphere at a temperature offrom about 700° to about 1200° F. The oxidizing atmosphere is suitablyair, although other gases comprising molecular oxygen may be employed.

Following the high temperature oxidation procedure the catalyst isusually reduced for a period of from about 1/2 to about 10 hours at atemperature in the range of from about 700° to about 1000° F. in thepresence of hydrogen. The catalyst may be used in a sulfided form. Thusafter reduction, the catalyst may be subjected to sulfidation by passinghydrogen sulfide, or other suitable sulfur containing compound, incontact therewith, preferably at an elevated temperature of from about500° to about 1100° F. The reduced catalyst is preferably sulfided bycontacting the catalyst with a stream of hydrogen containing from about1 to 20 percent or more by volume of hydrogen sulfide at elevatedtemperatures of from about 500° to about 1100° F. When the petroleumhydrocarbon to be hydrocracked contains sulfur compounds, by design orotherwise, sulfidation may be suitably effected in situ in the initialstages of the hydrocracking process.

The catalyst composite, prepared in accordance with the method of thisinvention, is preferably employed in a reaction zone as a fixed bed. Thehydrocarbon charge stock after being combined with hydrogen in an amountof from about 2000 to about 20,000 standard cubic feet per barrel, andpreferably at least about 5000 standard cubic feet per barrel, isintroduced into the reaction zone. The charge stock may be in a liquid,vapor, or liquid-vapor phase mixture, depending upon the temperature,pressure, proportion of hydrogen and the boiling range of the chargestock being processed. The liquid hourly space velocity through thereaction zone will be in excess of about 0.2 and generally in the rangeof from about 1.0 to about 15.0. The source of hydrogen being admixedwith a hydrocarbon charge stock may comprise a hydrogen-rich gas streamwhich is withdrawn from a high-pressure, low-temperature separation zoneand recycled to supply at least a portion of such hydrogen. Excesshydrogen resulting from the various dehydrogenation reactions effectedin a catalytic reforming unit may also be employed in admixture with thehydrocarbon charge. The reaction zone will operate under an imposedpressure within the range of from about 80 to about 3000 pounds persquare inch gauge. The catalyst bed inlet temperature is maintainedwithin the range of from about 350° to about 800° F. Since thehydrocracking reactions are exothermic, the outlet temperature or thetemperature at the bottom of the catalyst bed will be significantlyhigher than that at the inlet thereto. The degree of exothermicityexhibited by the temperature rise across the catalyst bed is at leastpartially dependent upon the character of the charge stock passingtherethrough, the rate at which the normally liquid hydrocarbon chargecontacts the catalyst bed, the intended degree of conversion tolower-boiling-hydrocarbon products, etc. In any event, the catalyst bedinlet temperature will be such that the exothermicity Of the reactionstaking place does not cause the temperature at the outlet of the bed toexceed about 900° F., and preferably 850° F. The operation may also beeffected as a moving-bed type, or suspensoid type of operation in whichthe catalyst, hydrocarbon and hydrogen are admixed and passed as aslurry through the reaction zone.

Although the method of preparing the catalyst, and careful selection ofoperating conditions within the ranges hereinbefore set forth, extendthe effective life of the catalyst composite, regeneration thereof mayeventually become desired due to the natural deterioration of thecatalytically active metallic components. The catalyst composite isreadily regenerated by treating the same in an oxidizing atmosphere, ata temperature of from about 750° to about 850° F., and burning coke andother heavy hydrocarbonaceous material therefrom. The catalyst compositemay then be subjected to the reducing action in hydrogen, in situ, at atemperature within the range of from about 1000° to about 1200° F. Ifdesirable, the catalyst may then be sulfided in the same manner as freshcatalyst as hereinbefore described.

Although not essential to successful hydroprocessing in all cases, infact detrimental in some, a halogen component may be incorporated intothe catalyst composite. Accordingly, a preferred catalyst composite, foruse in the present process, comprises a combination of a nickelcomponent and a tungsten component with a silica-alumina carriermaterial wherein said carrier material is co-gelled silica-aluminaconsisting of from about 43 percent to about 57 percent by weightalumina and from about 57 percent to 43 percent by weight silica andwherein said components are present in amounts sufficient to result inthe composite containing, on an elemental basis, about 2 to about 10percent by weight of the nickel component and about 8 to about 20percent by weight of the tungsten component and a halogen component.Although the precise form of the chemistry of the association of thehalogen component with the carrier material and metallic components isnot accurately known, it is customary in the art to refer to the halogencomponent as being combined with the carrier material, or with the otheringredients of the catalyst. The combined halogen may be eitherfluorine, chlorine, iodine, bromine, or mixtures thereof. Of thesefluorine and particularly chlorine are preferred for the hydrocarbonhydroprocesses encompassed by the present invention. The halogen may beadded to the carrier material in any suitable manner, and either duringpreparation of the carrier or before, or after the addition of the othercomponents. For example, the halogen may be added at any stage of thepreparation of the carrier material, or to the calcined carriermaterial, as an aqueous solution of an acid such as hydrogen fluoride,hydrogen chloride, hydrogen bromide, hydrogen iodide, etc. The halogencomponent or a portion thereof may be composited with the carriermaterial during the impregnation of the latter with the metalcomponents. The inorganic oxide hydrosol, which is typically utilized toform an amorphous carrier material, may contain halogen and thuscontribute at least a portion of the halogen component to the finalcomposite. The quantity of halogen is such that the final catalyticcomposite contains about 0.1% to about 1.5% by weight, and preferablyfrom about 0.5% to about 1.2%, calculated on an elemental basis.

In embodiments of the present invention wherein the instant catalystcomposite is used for the hydrogenation of hydrogenatable hydrocarbons,it is ordinarily a preferred practice to include an alkaline earth metalcomponent in the composite. More precisely, this optional component isselected from the group consisting of the compounds of the alkali metals-- cesium, rubidium, potassium, sodium, and lithium -- and the compoundsof the alkaline earth metals -- calcium, strontium, barium andmagnesium. Generally, good results are obtained in these embodimentswhen this component constitutes about 1 to about 5 weight percent of thecomposite, calculated on an elemental basis. This optional alkali oralkaline earth metal component can be incorporated in the composite inany of the known ways, with impregnation with an aqueous solution of asuitable water-soluble, decomposable compound being preferred.

An optional ingredient for the catalyst of the present invention is aFriedel-Crafts metal halide component. This ingredient is particularlyuseful in hydrocarbon conversion embodiments of the present inventionwherein it is preferred that the catalyst utilized has a strong acid orcracking function associated therewith -- for example, an embodimentwherein hydrocarbons are to be hydrocracked or isomerized with thecatalyst of the present invention. Suitable metal halides of theFreidel-Crafts type include aluminum chloride, aluminum bromide, ferricchloride, ferric bromide, zinc chloride and the like compounds, with thealuminum halides and particularly aluminum chloride ordinarily yieldingbest results. Generally, this optional ingredient can be incorporatedinto the composite of the present invention by any of the conventionalmethods for adding metallic halides of this type; however, best resultsare ordinarily obtained when the metallic halide is sublimed onto thesurface of the carrier material according to the preferred methoddisclosed in U.S. Pat. No. 2,999,074. The component can generally beutilized in any amount which is catalytically effective, with a valueselected from the range of about 1 to about 100 wt. % of the carriermaterial generally being preferred. When used in many of thehydrogen-consuming processes hereinbefore described, the foregoingquantities of metallic components will be combined with a carriermaterial of alumina and silica, wherein the silica concentration is10.0% to about 90.0% by weight.

Regardless of the details of how the components of the catalyst arecombined with the porous carrier material, the final catalyst generallywill be dried at a temperature of from about 200° to about 600° F. for aperiod of at least about 2 to about 24 hours or more, and finallycalcined or oxidized at a temperature of about 700° to about 1100° F. inan air atmosphere for a period of about 0.5 to about 10 hours. Because ahalogen component may be utilized in the catalyst, best results aregenerally obtained when the halogen content of the catalyst is adjustedduring the calcination step by including a halogen or ahalogen-containing compound in the air atmosphere utilized. Inparticular, when the halogen component of the catalyst is chlorine, itis preferred to use a mole ratio of H₂ O to HCl of about 5:1 to about100:1 during at least a portion of the calcination step in order toadjust the final chlorine content of the catalyst to a range of about0.5 to about 1.5 weight percent.

The resulting reduced catalytic composite may, in some cases, bebeneficially subjected to a presulfiding operation designed toincorporate in the catalytic composite from about 0.05 to about 0.5weight percent sulfur calculated on an elemental basis. Preferably, thispresulfiding treatment takes place in the presence of hydrogen and asuitable sulfur-containing compound such as hydrogen sulfide, lowermolecular weight mercaptans, organic sulfides, etc. Typically, thisprocedure comprises treating the selectively reduced catalyst with asulfiding gas such as a mixture of hydrogen and hydrogen sulfide havingabout 10 moles of hydrogen per mole of hydrogen sulfide at conditionssufficient to effect the desired incorporation of sulfur, generallyincluding a temperature ranging from about 50° up to about 1100° F. ormore. It is generally a good practice to perform this presulfiding stepunder substantially water-free conditions.

According to the present invention, a hydrocarbon charge stock andhydrogen are contacted with a catalyst of the type described above in ahydrocarbon conversion zone. This contacting may be accomplished byusing the catalyst in a fixed bed system, a moving bed system, afluidized bed system, or in a batch type operation; however, in view ofthe danger of attrition losses of the valuable catalyst and of wellknown operational advantages, it is preferred to use a fixed bed system.In this system, a hydrogen-rich gas and the charge stock are preheatedby any suitable heating means to the desired reaction temperature andthen are passed, into a conversion zone containing a fixed bed of thecatalyst type previously characterized. It is, of course, understoodthat the conversion zone may be one or more separate reactors withsuitable means therebetween to insure that the desired conversiontemperature is maintained at the entrance to each reactor. It is alsoimportant to note that the reactants may be contacted with the catalystbed in either upward, downward, or radial flow fashion with the latterbeing preferred. In addition, the reactants may be in the liquid phase,a mixed liquid-vapor phase, or a vapor phase when they contact thecatalyst.

The operating conditions imposed upon the reaction zone are dependentupon the particular hydroprocessing being effected. However, theseoperating conditions will include a pressure from about 400 to about5000 psig., a liquid hourly space velocity of about 0.1 to about 10.0,and a hydrogen circulation rate within the range of about 1,000 to about50,000 standard cubic feet per barrel. In view of the fact that thereactions being effected are exothermic in nature, an increasingtemperature gradient is experienced as the hydrogen and feed stocktraverses the catalyst bed. For any given hydrogen-consuming process, itis desirable to maintain the maximum catalyst bed temperature belowabout 900° F., which temperature is virtually identical to thatconveniently measured at the outlet of the reaction zone.Hydrogen-consuming processes are conducted at a temperature in the rangeof about 200° to about about 900° F., and it is intended herein that thestated temperature of operation alludes to the maximum catalyst bedtemperature. In order to assure that the catalyst bed temperature doesnot exceed the maximum allowed for a given process, the use ofconventional quench streams, either normally liquid or gaseous,introduced at one or more intermediate loci of the catalyst bed, may beutilized. In some of the hydrocarbon hydroprocesses encompassed by thepresent invention, and especially where hydrocracking a heavyhydrocarbonaceous material to produce lower-boiling hydrocarbonproducts, that portion of the normally liquid product effluent boilingabove the end point of the desired product will be recycled to combinewith the fresh hydrocarbon charge stock. In these situations, thecombined liquid feed ratio (defined as volumes of total liquid charge tothe reaction zone per volume of fresh feed charge to the reaction zone)will be within the range of about 1.1 to about 6.0.

The drawing included in the instant application is for the purpose ofvisually demonstrating the improvements and advantages afforded by themanufacture of silica-alumina-nickel-tungsten hydrocracking catalystaccording to the present invention.

The following example is presented in illustration of the catalyst ofthis invention and a method of preparation thereof, and is not intendedas an undue limitation on the generally broad scope of the invention asset out in the appended claims.

EXAMPLE I

This example described the preparation and testing of threesilica-alumina-nickel-tungsten catalysts each of which contains 8%nickel and 18% tungsten and which contain 40, 50 and 63 weight percentalumina, respectively. The co-gelled silica-alumina support material foreach catalyst was prepared by the hereinabove described oil-drop methodand the ratio of silica and alumina sources was selected to yield afinished support material which had the desired alumina content. Thefinished support material was in the form of 1/16 inch spheres.

A batch of co-gelled support material containing 40 weight percentalumina was impregnated with an aqueous solution of nickel nitrate andammonium metatungstate. The impregnated spheres were dried and thenoxidized at a temperature of 1100° F. The concentration of the metalsalts in the aqueous impregnation solution was selected to yield afinished catalyst which contained 8 weight percent nickel and 18 weightpercent tungsten. This batch of finished hydrocracking catalyst willhereinafter be referred to as catalyst 1.

A batch of co-gelled support material containing 50 weight percentalumina was then used to prepare catalyst 2 in the same manner ascatalyst 1. Catalyst 2 also contained 8 weight percent nickel and 18weight percent tungsten.

A batch of co-gelled support material which contained 63 weight percentalumina was impregnated to yield a finished catalyst 3 containing 8weight percent nickel and 18 weight percent tungsten in exactly the samemanner as the two previous preparations.

Each of the catalysts prepared in this manner were then used in thehydrocracking of a light vacuum gas oil whose properties are summarizedin Table I.

                  TABLE I:                                                        ______________________________________                                        Properties of Light Vacuum Gas Oil                                            ______________________________________                                        API ° Gravity at 60° F.                                                                  36.7                                                 Specific Gravity at 60° F.                                                                      0.8413                                               Distillation, ° F.                                                     IBP                      550                                                  10                       635                                                  30                       688                                                  50                       716                                                  70                       742                                                  90                       785                                                  E.P.                     856                                                  Total Sulfur, wt. %      0.07                                                 Total Nitrogen, wt. %    0.044                                                Aromatics, vol. %        12.7                                                 Paraffins and Naphthenes, vol. %                                                                       87.3                                                 Pour Point, ° F.  80                                                   ______________________________________                                    

In each case, the light vacuum gas oil was processed with a reactorpressure of 2000 psig., a liquid hourly space velocity of 1.0, ahydrogen circulation rate of 9500 scf./bbl. and at a catalyst bedtemperature which was required to yield a 315° + product with a -5° F.pour point. Catalysts 1, 2 and 3 required catalyst temperatures of 763°,754° F. and 779° F., respectively, to yield the desired productcharacteristics. These data are presented in tabular form in Table IIand in graphical form in the accompanying drawing.

                  TABLE II:                                                       ______________________________________                                        Evaluation For Hydrocracking Activity                                         Catalyst Identity 1       2       3                                           ______________________________________                                        Alumina Concentration                                                                           40      50      63                                          Reactor Temperature Required                                                  For -5° F. Pour Point                                                                    763     754     779                                         ______________________________________                                    

From the data presented in foregoing Table I and with reference to theaccompanying drawing, it will be seen that the three catalysts'increasing concentrations of alumina in the carrier material, the latterranging from 40% to 63% by weight, did not produce normally liquidhydrocarbon products with improved pour point characteristics at thelowest catalyst bed temperature. This is clearly brought out uponcomparing the results obtained through the use of catalysts 1, 2 and 3which results in catalyst bed temperatures of 763°, 754° and 779° F.,respectively, for the desired product characteristics. Datum points 1, 2and 3 in the drawing are representative of the results obtained withcatalysts 1, 2 and 3 respectively. These data were employed in preparingcurve 4 of the drawing, which curve clearly illustrates the criticalityattached to an alumina concentration within the range of about 43% toabout 57% by weight, in order to produce a liquid product with thedesired characteristics at the lowest catalyst bed temperature. Theadditional economic advantages afforded through this particular resultwill be readily recognized by those possessing skill within the art ofpetroleum refining processes.

Specific operating conditions, processing techniques, particularcatalytic composites and other individual process details will be givenin the following detailed description of several of the hydrocarbonhydroprocesses to which the present invention is applicable. These willbe presented by way of examples given in conjunction withcommercially-scaled operating units. In presenting these examples, it isnot intended that the invention be limited to the specificillustrations, nor is it intended that a given process be limited to theparticular operating conditions, catalytic composite, processingtechniques, charge stocks, etc. It is understood, therefore, that thepresent invention is merely illustrated by the specifics hereinafter setforth.

EXAMPLE II

In this example, the present invention is illustrated as applied to thehydrogenation of aromatic hydrocarbons such as benzene, toluene, thevarious xylenes, naphthalenes, etc., to form the corresponding cyclicparaffins. When applied to the hydrogenation of aromatic hydrocarbons,which are contaminated by sulfurous compounds, primarily thipheniccompounds, the process is advantageous in that it affords 100.0%conversion without the necessity for the substantially complete priorremoval of the sulfur compounds. The corresponding cyclic paraffins,resulting from the hydrogenation of the aromatic nuclei, includecompounds such as cyclohexane, mono-, di-, tri-substituted cyclohexanes,decahydronaphthalene, tetrahydronaphthalene, etc., which find widespreaduse as a variety of commercial industries in the manufacture of nylon,as solvents for various fats, oils, waxes, etc.

Aromatic concentrates are obtained by a multiplicity of techniques. Forexample, a benzene-containing fraction may be subjected to distillationto provide a heart-cut which contains the benzene. This is thensubjected to a solvent extraction process which separates the benzenefrom the normal or iso-paraffinic components, and the naphthalenescontained therein. Benzene is readily recovered from the selectedsolvent by way of distillation, and in a purity of 99.0% or more.Heretofore, the hydrogenation of aromatic hydrocarbons, for examplebenzene, has been effected with a nickel-containing catalyst. This isextremely disadvantageous in many respects, and especially from thestandpoint that nickel is quite sensitive to the minor quantity ofsulfurous compounds which may be contained in the benzene concentrate.In accordance with the present process, the benzene is hydrogenated incontact with a catalytic composite comprising a 50/50 silica-aluminacarrier material and containing 8 weight percent nickel and 18 weightpercent tungsten. Operating conditions include a maximum catalyst bedtemperature in the range of about 200° to about 800° F., a pressure offrom 500 to about 2,000 psig., a liquid hourly space velocity of about1.0 to about 10.0 and a hydrogen circulation rate in an amountsufficient to yield a mole ratio of hydrogen to cyclohexane, in theproduct effluent from the last reaction zone, not substantially lessthan about 4.0:1. Although not essential, one preferred operatingtechnique involves the use of three reaction zones, each of whichcontains approximately one-third of the total quantity of catalystemployed. The process is further facilitated when the total freshbenzene is added in three approximately equal portions, one each to theinlet of each of the three reaction zones.

The hydrogenation process will be described in connection with acommercially-scaled unit having a total fresh benzene feed capacity ofabout 1,488 barrels per day. Makeup gas in an amount of about 741.6mols/hr. is admixed with 2,396 Bbl./day (about 329 mols/hr.) of acyclohexane recycle stream, the mixture being at a temperature of about137° F., and further mixed with 96.24 mols/hr. (582 Bbl./day) of thebenzene feed; the final mixture constitutes the total charge to thefirst reaction zone.

Following suitable heat-exchange with various hot effluent streams, thetotal feed to the first reaction zone is at a temperature of 385° F. anda pressure of 460 psig. The reaction zone effluent is at a temperatureof 606° F. and a pressure of about 450 psig. The total effluent from thefirst reaction zone is utilized as a heat-exchange medium, in a steamgenerator, whereby the temperature is reduced to a level of about 545°F. The cooled effluent is admixed with about 98.5 moles per hour (596Bbl./day) of fresh benzene feed, at a temperature of 100° F.; theresulting temperature is 400° F., and the mixture enters the secondreaction zone at a pressure of about 440 psig. The second reaction zoneeffluent, at a pressure of 425 psig. and a temperature of 611° F., isadmixed with 51.21 mols/hr. (310 Bbl./day) of fresh benzene feed, theresulting mixture being at a temperature of 578° F. Following its use asa heat-exchange medium, the temperature is reduced to 400° F., and themixture enters the third reaction zone at a pressure of 415 psig. Thethird reaction zone effluent is at a temperature of about 500° F. and apressure of about 400 psig. Through utilization as a heat-exchangemedium, the temperature is reduced to a level of about 244° F., andsubsequently reduced to a level of about 115° F. by use of an air-cooledcondenser. The cooled third reaction zone effluent is introduced into ahigh pressure separator, at a pressure of about 370 psig.

A hydrogen-rich vaporous phase is withdrawn from the high pressureseparator and recycled by way of compressive means, at a pressure ofabout 475 psig., to the inlet of the first reaction zone. A portion ofthe normally liquid phase is recycled to the first reaction zone as thecyclohexane concentrate hereinbefore described. The remainder of thenormally liquid phase is passed into a stabilizing column functioning atan operating pressure of about 250 psig., a top temperature of about160° and a bottom temperature of about 430° F. The cyclohexane productis withdrawn from the stabilizer as a bottoms stream, the overheadstream being vented to fuel. The cyclohexane concentrate is recovered inan amount of about 245.80 moles per hour, of which only about 0.60 molesper hour constitutes other hexanes. In brief summation, of the 19,207pounds per hour of fresh benzene feed, 20,685 pounds per hour ofcyclohexane product is recovered.

EXAMPLE III

Another hydrocarbon hydroprocessing scheme, to which the presentinvention is applicable, involves the hydrorefining of coke-forminghydrocarbon distillates. The hydrocarbon distillates are generallysulfurous in nature, and contain mono-olefinic, di-olefinic and aromatichydrocarbons. Through the utilization of a catalytic composite preparedaccording to the present invention increased selectivity and stabilityof operation is obtained; selectivity is most noticeable with respect tothe retention of aromatics, and in hydrogenating conjugated di-olefinicand mono-olefinic hydrocarbons. Such charge stocks generally result fromdiverse conversion processes including the catalytic and/or thermalcracking of petroleum, sometimes referred to as pyrolysis, thedestructive distillation of wood or coal, shale or retorting, etc. Theimpurities in these distillate fractions must necessarily be removedbefore the distillates are suitable for their intended use, or whichwhen removed, enhance the value of the distillate fraction for furtherprocessing. Frequently, it is intended that these charge stocks besubstantially desulfurized, saturated to the extent necessary to removethe conjugated di-olefins, while simultaneously retaining the aromatichydrocarbons. When subjected to hydrorefining for the purpose ofremoving the contaminating influences, there is encountered difficultyin effecting the desired degree of reaction due to the formation of cokeand other carbonaceous material.

As utilized herein, "hydrogenating" is intended to be synonymous with"hydrorefining". The purpose is to provide a highly selective and stableprocess for hydrogenating coke-forming hydrocarbon distillates, and thisis accomplished through the use of a fixed-bed catalytic reaction systemutilizing a catalyst of the present invention. There exists twoseparate, desirable routes for the treatment of coke-formingdistillates, for example a pyrolysis naphtha by-product. One such routeis directed toward a product suitable for use in certain gasolineblending. With this as the desired object, the process can be effectedin a single stage, or reaction zone, with the catalytic compositehereinafter specifically described as the first-stage catalyst. Theattainable selectively in this instance resides primarily in thehydrogenation of highly reactive double bonds. In the case of conjugateddi-olefins, the selectively afforded restricts the hydrogenation toproduce mono-olefins, and, with respect to the styrenes, for example,the hydrogenation is inhibited to produce alkyl benzenes without "ring"saturation. The selectivity is accomplished with a minimum of polymerformation either to "gums", or lower molecular weight polymers whichwould necessitate a re-running of the product before blending togasoline would be feasible. Other advantages of restricting thehydrogenating of the conjugated di-olefins, such as 1,5 normal hexadieneare not usually offensive in suitably inhibited gasolines in somelocales, and will not react in this first stage. Some fresh chargestocks are sufficiently low in mercaptan sulfur content that directgasoline blending may be considered, although a mild treatment formercaptan sulfur removal might be necessary. These considerations aregenerally applicable to foreign markets, particularly European, whereolefinic and sulfur-containing gasolines are not too objectionable. Itmust be noted that the sulfurous compounds, and the monoolefins, whethervirgin, or products of di-olefin partial saturation, are unchanged inthe single, or first-stage reaction zone. Where however the desired endresult is aromatic hydrocarbon retention, intended for subsequentextraction, the two-stage route is required. The mono-olefins must besubstantially saturated in the second stage to facilitate aromaticextraction by way of currently utilized methods. Thus, the desirednecessary hydrogenation involves saturation of the mono-olefins, as wellas sulfur removal, the latter required for an acceptable ultimatearomatic product. Attendant upon this is the necessity to avoid evenpartial saturation of aromatic nuclei.

With respect to one catalytic composite, its principal function involvesthe selective hydrogenation of conjugated di-olefinic hydrocarbons tomono-olefinic hydrocarbons. The particular catalytic composite possessesunusual stability notwithstanding the presence of relatively largequantities of sulfurous compounds in the fresh charge stock. Thecatalytic composite comprises a 50/50 silica-alumina carrier materialand containing 8 weight percent nickel, 18 weight percent tungsten andan alkali-metal component, the latter being preferably potassium and/orlithium. It is especially preferred, for use in this particularhydrocarbon hydroprocessing scheme, that the catalytic composite besubstantially free from any "acid-acting" propensities. The catalyticcomposite, utilized in the second reaction zone for the primary purposeof effecting the destructive conversion of sulfurous compounds intohydrogen sulfide and hydrocarbons, is a composite similar to the firstreaction zone catalyst without an alkali metal component. Through theutilization of a particular sequence of processing steps, and the use ofthe foregoing described catalyst composites, the formation of highmolecular weight polymers and co-polymers is inhibited to a degree whichpermits processing for an extended period of time. Briefly, this isaccomplished by initiating the hydrorefining reactions at temperaturesbelow about 500° F., at which temperatures the coke-forming reactionsare not promoted. The operating conditions within the second reactionzone are such that the sulfurous compounds are removed without incurringthe detrimental polymerization reactions otherwise resulting were it notfor the saturation of the conjugated di-olefinic hydrocarbons within thefirst reaction zone.

The hydrocarbon distillate charge stock, for example, a light naphthaby-product from a commercial cracking unit designed and operated for theproduction of ethylene, having a gravity of about 34.0° API, a brominenumber of about 35.0, a diene value of about 17.5 and containing about1,600 ppm. by weight of sulfur and 75.9 vol.% aromatic hydrocarbons, isadmixed with recycled hydrogen. This light naphtha co-product has aninitial boiling point of about 164° F. and an end boiling point of about333° F. The hydrogen circulation rate is within the range of from about1,000 to about 10,000 scf./bbl., and preferably in the narrower range offrom 1,500 to about 6,000 scf./bbl. The charge stock is heated to atemperature such that the maximum catalyst temperature is in the rangeof from about 200° to about 500° F., by way of heat-exchange withvarious product effluent streams, and is introduced into the firstreaction zone at an LHSV in the range of about 0.5 to about 10.0. Thereaction zone is maintained at a pressure of from 400 to about 1,000psig., and preferably at a level in the range of from 500 psig. to about900 psig.

The temperature of the product effluent from the first reaction zone isincreased to a level above about 500° F., and preferably to result in amaximum catalyst temperature in the range of 600° to 900° F. When theprocess is functioning efficiently, the diene value of the liquid chargeentering the second catalytic reaction zone is less than about 1.0 andoften less than about 0.3. The conversion of nitrogenous and sulfurouscompounds, and the saturation of mono-olefins, contained within thefirst zone effluent, is effected in the second zone. The secondcatalytic reaction zone is maintained under an imposed pressure of fromabout 400 to about 1,000 psig., and preferably at a level of from about500 to about 900 psig. The two-stage process is facilitated when thefocal point for pressure control is the high pressure separator servingto separate the product effluent from the second catalytic reactionzone. It will, therefore, be maintained at a pressure slightly less thanthe first catalytic reaction zone, as a result of fluid flow through thesystem. The LHSV through the second reaction zone is about 0.5 to about10.0, based upon fresh feed only. The hydrogen circulation rate will bein a range of from 1,000 to about 10,000 scf./bbl., and preferably fromabout 1,000 to about 8,000 scf./bbl. Series-flow through the entiresystem is facilitated when the recycle hydrogen is admixed with thefresh hydrocarbon charge stock. Make-up hydrogen, to supplant thatconsumed in the overall process, may be introduced from any suitableexternal source, but is preferably introduced into the system by way ofthe effluent line from the first catalytic reaction zone to the secondcatalytic reaction zone.

With respect to the naphtha boiling range portion of the producteffluent, the sulfur concentration is about 0.1 ppm., the aromaticconcentration is about 75.1% by volume, the bromine number is less thanabout 0.3 and the diene value is essentially "nil".

With charge stocks having exceedingly high diene values, a recyclediluent is employed in order to prevent an excessive temperature rise inthe reaction system. Where so utilized, the source of the diluent ispreferably a portion of the normal liquid product effluent from thesecond catalytic reaction zone. The precise quantity of recycle materialvaries from feed stock to feed stock; however, the rate at any giventime is controlled by monitoring the diene value of the combined liquidfeed to the first reaction zone. As the diene value exceeds a level ofabout 25.0, the quantity of recycle is increased, thereby increasing thecombined liquid feed ratio; when the diene value approaches a level ofabout 20.0, or less, the quantity of recycle diluent may be lessened,thereby decreasing the combined liquid feed ratio.

With another so-called pyrolysis gasoline, having a gravity of about36.4° API, containing 600 ppm. by weight of sulfur, 78.5% by volume ofaromatics, and having a bromine number of 45 and a diene value of 25.5it is initially processed in a first reaction zone containing acatalytic composite comprising a 50/50 silica-alumina carrier materialand containing 8 weight percent nickel, 18 weight percent tungsten and0.5 weight percent lithium. The fresh feed charge rate is 3,300bbl./day, and this is admixed with 2,475 bbl./day of the normally liquiddiluent. Based upon fresh feed only, the LHSV is 2.5 and the hydrogencirculation rate is 1,750 scf./bbl. The charge is raised to atemperature of about 250° F., and enters the first reaction zone at apressure of about 850 psig. The product effluent emanates from the firstreaction zone at a pressure of about 830 psig. and a temperature ofabout 350° F. The effluent is admixed with about 660 scf./bbl. ofmake-up hydrogen, and the temperature is increased to a level of about545° F., the heated stream is introduced into the second reaction zoneunder a pressure of about 790 psig. The LHSV, exclusive of the recyclediluent, is 2.5, and the hydrogen circulation rate is about 1,500. Thesecond reaction zone contains a catalyst comprising a 50/50silica-alumina carrier material and containing 8 weight percent nickeland 18 weight percent tungsten. The reaction product effluent isintroduced following its use as a heat-exchange medium and furthercooling, to reduce its temperature from 620° to a level of 100° F., intoa high-pressure separator at a pressure of about 750 psig. The normallyliquid stream from the cold separator is introduced into a reboiledstripping column for hydrogen sulfide removal and depentanization. Thehydrogen sulfide stripping column functions at conditions of temperatureand pressure required to concentrate a C₆ to C₉ aromatic stream as abottoms fraction. With respect to the overall product distribution, only690 lbs/hr of pentanes and lighter hydrocarbons is indicated in thestripper overhead. The aromatic concentrate is recovered in an amount ofabout 40,070 lbs/hr (the fresh feed is 40,120 lbs/hr); these results areachieved with a hydrogen consumption of only 660 scf./bbl. With respectto the desired product, the aromatic concentration 78.0, the sulfurconcentration is less than 1.0 ppm. by weight, and the diene value isessentially "nil".

EXAMPLE IV

This example is presented to illustrate still another hydrocarbonhydroprocessing scheme for the improvement of the jet fuelcharacteristics of sulfurous kerosene boiling range fractions. Theimprovement is especially noticeable in the IPT Smoke Point, theconcentration of aromatic hydrocarbons and the concentration of sulfur.A two-stage process wherein desulfurization is effected in the firstreaction zone at relatively mild severities which result in a normallyliquid product effluent containing from about 15 to about 35 ppm. ofsulfur by weight. Aromatic saturation is the principal reaction effectedin the second reaction zone, having disposed therein a catalyticcomposite comprising a 50/50 silica-alumina carrier material andcontaining 8 weight percent nickel, 18 weight percent tungsten, and 0.6weight percent combined chloride.

Suitable charge stocks are kerosene fractions having an initial boilingpoint as low as about 300° F., and an end boiling point as high as about575°, and, in some instances, up to 600° F. Examplary of such kerosenefractions are those boiling from about 300° to about 550° F., from 330°to about 500°, from 330° to about 530° F., etc. As a specific example, akerosene obtained from hydrocracking a Mid-continent slurry oil, havinga gravity of about 30.5° API, an initial boiling point of about 388° F.,an end boiling point of about 522° F., has an IPT Smoke Point of 17.1mm., and contains 530 ppm. of sulfur and 24.8% by volume of aromatichydrocarbons. Through the use of the catalytic process of the presentinvention, the improvement in the jet fuel quality of such a kerosenefraction is most significant with respect to raising the IPT SmokePoint, and reducing the concentration of sulfur and the quantity ofaromatic hydrocarbons. Specifications regarding the poorest quality ofjet fuel, commonly referred to as Jet A, Jet-A1 and Jet-B call for asulfur concentration of about 0.3% by weight maximum (3,000 ppm.), aminimum IPT Smoke Point of 25 mm. and a maximum aromatic concentrationof about 20.0 vol.%.

The charge stock is admixed with circulating hydrogen in an amountwithin the range of from about 1,000 to about 2,000 scf./bbl. Thismixture is heated to a temperature level necessary to control themaximum catalyst bed temperature below about 750° F., and preferably notabove 700°, with a lower catalyst bed temperature of about 600° F. Thecatalyst, a well known standard desulfurization catalyst containingabout 2.2% by weight of cobalt and about 5.7% by weight of molybdenum,composited with alumina is disposed in a reaction zone maintained underan imposed pressure in the range of from about 500 to about 1,000 psig.The LHSV is in the range of about 0.5 to about 10.0, and preferably fromabout 0.5 to about 5.0. The total product effluent from this firstcatalytic reaction zone is separated to provide a hydrogen-rich gaseousphase and a normally liquid hydrocarbon stream containing 15 ppm. toabout 35 ppm. of sulfur by weight. The normally liquid phase portion ofthe first reaction zone effluent is utilized as the fresh feed chargestock to the second reaction zone. In this particular instance, thefirst reaction zone decreases the sulfur concentration of about 25 ppm.,the aromatic concentration to about 16.3% by volume, and has increasedthe IPT Smoke Point to a level of about 21.5 mm.

The catalytic composite within the second reaction zone is describedhereinabove. The reaction zone is maintained at a pressure of about 400to about 1,500 psig., and the hydrogen circulation rate is in the rangeof 1,500 to about 10,000 scf./bbl. The LHSV, hereinbefore defined, is inthe range of from about 0.5 to about 5.0, and preferably from about 0.5to about 3.0. It is preferred to limit the catalyst bed temperature inthe second reaction zone to a maximum level of about 750° F. With acatalyst of this particular chemical and physical characteristics,optimum aromatic saturation, processing a feed stock containing fromabout 15 to about 35 ppm. of sulfur, is effected at maximum catalyst bedtemperatures in the range of about 625° to about 750° F. With respect tothe normally liquid kerosene fraction, recovered from the condensedliquid removed from the total product effluent, the sulfur concentrationis effectively "nil", being about 0.1 ppm. The quantity of aromatichydrocarbons has been decreased to a level of about 0.75% by volume, andthe IPT Smoke Point has been increased to about 36.3 mm.

EXAMPLE V

This illustration of a hydrocarbon hydroprocessing scheme, encompassedby my invention is one which involves hydrocracking heavyhydrocarbonaceous material into lower boiling hydrocarbon products. Inthis instance, the preferred catalysts contain a 50/50 silica-aluminacarrier material, 8 weight percent nickel, 18 weight percent tungstenand 0.7 weight percent combined chloride.

Most of the virgin stocks, intended for hydrocracking, are contaminatedby sulfurous compounds and nitrogenous compounds, and, in the case ofthe heavier charge stocks, various metallic contaminants, insolubleasphalts, etc. Contaminated charge stocks are generally hydrorefined inorder to prepare a charge suitable for hydrocracking. Thus, thecatalytic process of the present invention can be beneficially utilizedas the second stage of a two-stage process, in the first stage of whichthe fresh feed is hydrorefined.

Hydrocracking reactions are generally effected at elevated pressures inthe range of about 800 to about 5,000 psig., and preferably at someintermediate level of 1,000 to about 3,500 psig. Liquid hourly spacevelocities of about 0.25 to about 10.0 will be suitable, the lower rangegenerally reserved for the heavier stocks. The hydrogen circulation ratewill be at least about 3,000 scf./bbl., with an upper limit of about50,000 scf./bbl., based upon fresh feed. For the majority of feedstocks, hydrogen circulation in the range of 5,000 to 20,000 scf./bbl.will suffice. With respect to the LHSV, it is based upon fresh feed,notwithstanding the use of recycle liquid providing a combined liquidfeed ratio in the range of about 1.25 to about 6.0. The operatingtemperature again alludes to the temperature of the catalyst within thereaction zone, and is in the range of about 400° to about 900° F. Sincethe principal reactions are exothermic in nature, the increasingtemperature gradient, experienced as the charge stock traverses thecatalyst bed, results in an outlet temperature higher than that at theinlet to the catalyst bed. The maximum catalyst temperature should notexceed 900°, and it is generally a preferred technique to limit thetemperature increase to 100° F. or less.

Other possible constituent of the catalyst is a halogen component,either fluorine, chlorine, iodine, bromine, or mixtures thereof. Ofthese, it is preferred to utilize a catalyst containing fluorine and/orchlorine. The halogen component will be composited with the carriermaterial in such a manner as results in a final composite containingabout 0.1% to about 1.5% by weight of halogen, calculated on anelemental basis.

A specific illustration of this hydrocarbon hydroprocessing techniqueinvolves the use of a catalytic composite hereinabove described.

This catalyst is intended for utilization in the conversion of 16,000bbl./day of a blend of light gas oils to produce maximum quantities of aheptane-400° F. gasoline boiling range fraction. The charge stock has agravity of 33.8° API, contains 0.19% by weight of sulfur (1,900 ppm.)and 67 ppm. by weight of nitrogen, and has an initial boiling point of369° F., a 50% volumetric distillation temperature of 494° F. and an endboiling point of 655° F. The charge stock is initially subjected to aclean-up operation at maximum catalyst temperature of 750° F., acombined feed ratio of 1.0, an LHSV of 2.41 with a hydrogen circulationrate of about 5000 scf./bbl. The pressure imposed upon the catalystwithin the reaction zone is about 1,500 psig. Since at least a portionof the blended gas oil charge stock will be converted into alower-boiling hydrocarbon product, the effluent from this clean-upreaction zone is separated to provide a normally liquid, 400° F.-pluscharge for the hydrocracking reaction zone. The pressure imposed uponthe second reaction zone is about 1,500 psig., and the hydrogencirculation rate is about 8,000 scf./bbl. The original quantity of freshfeed to the clean-up reaction zone is about 16,000 bbl./day; followingseparation of the first zone effluent to provide the 400° F.-plus chargeto the second reaction zone, the charge to the second reaction zone isin an amount of about 12,150 bbl./day, providing an LHSV of 0.85. Thetemperature at the inlet to the catalyst bed is 665° F., and aconventional hydrogen quench stream is utilized to maintain the maximumreactor outlet temperature at about 700° F. Following separation of theproduct effluent from the second reaction zone, to concentrate thedesired gasoline boiling range fraction, the remaining 400° F.-plusnormally liquid material, in an amount of 7,290 bbl./day, is recycled tothe inlet of the second reaction zone, thus providing a combined liquidfeed ratio thereto of about 1.60. In the following table, there isindicated the product yield and distribution of this process. Withrespect to normally liquid hydrocarbons, for convenience includingbutanes, the yields are given in vol. %; with respect to the normallygaseous hydrocarbons, ammonia and hydrogen sulfide, the yields are givenin terms of wt. %. With respect to the first reaction zone, the hydrogenconsumption is 1.31% by weight of the fresh feed (741 scf./bbl.), andfor the hydrocracking reaction zone, 1.26% by weight of the fresh feedcharge stock, or 713 scf./bbl.

                  TABLE:                                                          ______________________________________                                        Hydrocracking Product Yield and Distribution                                  Component  Stage I  Stage II  Total                                           ______________________________________                                        Ammonia    0.01     --        0.01                                            Hydrogen   0.21     --        0.21                                            Sulfide                                                                       Methane    0.12     0.02      0.14                                            Ethane     0.22     0.40      0.62                                            Propane    1.03     3.48      4.51                                            Butanes    3.90     14.66     18.56                                           Pentanes   3.04     11.28     14.32                                           Hexanes    3.00     11.21     14.21                                           C.sub.7 -400° F.                                                                  18.85    49.56     68.41                                           400° F.-plus                                                                      75.92*   --        --                                              ______________________________________                                         *Charge to Stage II                                                      

With respect to both the butanes product and pentane product, the formeris indicated as being about 68.0% isobutane, while the latterconstitutes about 93.0% isopentanes. An analysis of the combinedpentane/hexane fraction indicates a gravity of 82.6° API, a clearresearch octane rating of 85.0 and a leaded research octane rating of99.0; it will be noted that this constitutes an excellent blendingcomponent for motor fuel. The desired heptane-400° F. product indicatesa gravity of 48.8° API, a clear research octane rating of 72.0 and aleaded research octane rating of 88.0. This gasoline boiling rangefraction constitutes about 34.0% by volume paraffins, 36.0% by volumenaphthenes and 30.0% by volume aromatic hydrocarbons. It will berecognized that this gasoline boiling range fraction constitutes anexcellent charge stock for a catalytic reforming unit to improve themotor fuel characteristics thereof.

The foregoing specification, and particularly the examples, indicatesthe methods by which the present invention is effected, and the benefitsafforded through the utilization thereof. pg,44

I claim as my invention:
 1. A process for hydrogenating a coke-forminghydrocarbon distillate containing conjugated di-olefinic andmono-olefinic hydrocarbons, and aromatics, which process comprisesreacting said distillate with hydrogen at a temperature below about 500°F. in contact with a catalytic composite comprising a combination of anickel component, and a tungsten component with a silica-alumina carriermaterial wherein said carrier material is co-gelled silica-aluminaconsisting of from about 43 percent to about 57 percent by weightalumina and from about 57 to about 43 percent by weight silica andwherein said components are present in amounts sufficient to result inthe composite containing, on an elemental basis, about 2 to about 10percent by weight of the nickel component and about 8 to about 20percent by weight of the tungsten component and an alkali metalcomponent, and recovering an aromatic/mono-olefinic hydrocarbonconcentrate substantially free from conjugated di-olefinic hydrocarbons.2. The process of claim 1 further characterized in that said carriermaterial consists of 50 percent by weight of alumina and wherein saidcomposite contains, on an elemental basis, about 7 to about 9 percent byweight nickel and about 17 to about 19 percent by weight of tungsten. 3.The process of claim 1 further characterized in that said conditionsinclude a pressure of from about 400 to about 5000 psig., a liquidhourly space velocity of from about 0.1 to about 10, a hydrogencirculation rate of from about 1000 to about 50,000 scf./Bbl. and amaximum catalyst temperature of from about 200° F. to about 900° F. 4.The process of claim 1 further characterized in that said catalyticcomposite contains from about 0.1% to about 1.5% by weight of a halogencomponent, on an elemental basis.
 5. The process of claim 1 furthercharacterized in that said catalytic composite contains from about 0.05to about 1 weight percent sulfur, calculated on an elemental basis.